Many conventional alkylation methods and processes are known and employed, varying more or less from one another, wherein isobutane is alkylated with olefins in the presence of sulphuric acids or other acid catalyst and an excess of isobutane. Several types of reaction vessels may be employed in these processes. The reaction steps of the several processes is typically cooled by indirect or direct heat exchange to control reaction temperature. Where the reaction step is heat exchanged, closed cycle refrigeration, effluent refrigeration or direct vaporization of reacting liquids may or may not be employed in a specific case.
In each of the typical and conventional alkylation processes, however, whatever the specific arrangements for conducting the reaction may be, or heat exchanging it, once the catalyst phase has been separated from the hydrocarbon phase of the reaction step effluent, the hydrocarbon component is typically passed to various stages of fractionation where alkylate product is separated from excess isoparaffinic hydrocarbons, in order that the latter may be recycled as feed to the reaction step to aid in the important goal and step of maintenance of a large proportional excess of isobutane in the reaction step.
The equipment involved in such fractionation separation typically includes a deisobutanizer tower of great expense and size. For example, an 8,000 barrel of product per day tower today has a cost in excess of $800,000. In view of such great cost it is eminently desirable to reduce the size of the deisobutanizer tower as much as possible. However, any change in this direction lies directly in the face of one of the most important functions of the fractionation system, specifically, to return as much isobutane as possible to the reaction step in order to maintain the optimum reaction conditions and produce the highest quality alkylate product.
Another important consideration in existing alkylation plants is how to increase the capacity of an existing plant by addition of one or more reaction vessels with a concomitant investment in as little additional fractionation equipment as possible. Once again, when the quantity of hydrocarbon phase effluent is increased, the requirement of separation of isoparaffinic hydrocarbons therefrom is also additionally increased. This generally means proportional addition of expensive fractionation equipment at great expense in time, cost, space and the like.
My U.S. Pat. No. 3,055,958 "Alkylation Effluent Flash Vaporization System", issued Sept. 25, 1962 shows effective means, apparatus and methods for meeting major aspects of the problems stated. Specifically, by use of my alkylation effluent flash vaporization process and apparatus prior to the fractionation steps, a substantial and important portion of the isoparaffinic hydrocarbons in the net hydrocarbon phase effluent from the reaction step is returned to the reaction step without reaching the fractionation stages. The methods disclosed in that Patent were adequate for their time with respect to obtaining the goals and objects related in that Patent and meeting the problems outlined above. However, in the present day and age, with the great cost of energy a reality for the present and foreseeable future, improved modes of carrying out alkylation effluent flash vaporization systems are called for. Specifically, the economics of steam as a heat exchanging medium available for use in such systems are very much in question. Accordingly, heat sources for use in the flash vaporization separation system must preferably be found of substantially different type to meet the needs of the present and future times.